Maximum reaction rate converter system for exothermic reactions

ABSTRACT

An ammonia converter system and method are disclosed. The reactor can alter the conversion of ammonia by controlling the reaction temperature of the exothermic reaction along the length of the reactor to parallel the equilibrium curve for the desired product. The reactor  100  can comprise a shell  101  and internal catalyst tubes  109.  The feed gas stream enters the reactor, flows through the shell  101,  and is heated by indirect heat exchange with the catalyst tubes  109.  The catalyst tubes  109  comprise reactive zones  122  having catalyst and reaction limited zones  124  that can comprise inert devices that function to both separate the reactive zones, increase heat transfer area, and reduce the temperature of the reaction mixture as the effluent passes through the catalyst tube  109.

BACKGROUND OF THE INVENTION

This invention relates to a converter for exothermic reactions, and moreparticularly to a converter and method such as for converting nitrogenand hydrogen to ammonia whereby reduced catalyst usage and/or greaterproduct yields are obtainable.

Ammonia is commonly manufactured by reacting nitrogen and hydrogen in asynthesis loop which can include a compressor, an ammonia synthesisreactor, and an ammonia condensation and recovery step. Unreactedsynthesis gas from the synthesis reaction is typically recycled from theammonia separator back to the compressor and reactor. The synthesis gascan contain argon, methane, and other inert components which aretypically removed as a purge stream, thereby avoiding buildup of inertsin the synthesis loop. The purge gas can be further processed in ahydrogen recovery unit, or alternatively supplied directly to the fuelsystem with or without additional treatment or hydrogen recovery.

Many ammonia production plants operate with a synthesis loop using aniron-based magnetite catalyst in the ammonia converters. Significantadvances in the manufacture of ammonia have included the use of highlyactive synthesis catalysts comprising a platinum group metal supportedon graphite-containing carbon, used alone or in conjunction with lessactive iron based catalysts, as described in U.S. Pat. Nos. 4,568,530,4,568,531, and 4,568,532. Desirably, the platinum group metal isruthenium, as more fully described in U.S. Pat. Nos. 4,122,040 and4,250,057. The highly active catalysts generally allow for increasedammonia production and/or the usage of smaller volumes of catalyst.

In general, contact of the reactants with a catalyst under suitabletemperature and pressure conditions effects an exothermic reaction. Theheating associated with exothermic reactions can have various positiveand negative effects on the reaction. Negative effects can include: poorproduction rates, deactivation of catalyst, production of unwantedby-products, and damage to the reaction vessel and piping. Mostcommonly, an excessive temperature increase in the reaction zone eitherlimits selectivity or reduces product yield.

Exothermic reaction processes can encompass a wide variety of feedstocksand products. Examples of moderately exothermic processes can includemethanol synthesis, ammonia synthesis, and the conversion of methanol toolefins. Examples of highly exothermic reactions can include oxidationreactions in general, phthalic anhydride manufacture by naphthalene ororthoxylene oxidation, acrylonitrile production from propane orpropylene, acrylic acid synthesis from acrolein, the conversion ofn-butane to maleic anhydride, the production of acetic acid by methanolcarbonylation, and methanol conversion to formaldehyde.

The efficiency of reversible exothermic reactions often depends on theability to remove the heat generated by the process. The reaction rateand equilibrium generally move oppositely with increasing temperature.Thus, higher reaction temperatures generally result in faster reactionrates and lower overall conversion, while lower reaction temperaturesgenerally result in slower reaction rates and higher overall conversion.For increased conversion in staged reversible exothermic reactions, ahigh temperature is employed in the early stages of the reaction wherethe reaction kinetics are more favorable. As the reaction progresses,the temperature in the later stages is reduced to take advantage of themore favorable equilibrium conditions. However, because the reaction isdone in stages with interstage cooling, the equilibrium and kinetics arerarely, or only for very briefly, balanced for the maximum reaction ratepossible. The present invention employs conditions approximating theoptimal reactor operating curve (or temperature progression) whichmaximizes the reaction rate along a path corresponding to a locus ofmaximum rates on a temperature-conversion plot. This type of plotgenerally follows a decreasing temperature profile moving from thereactor inlet to outlet.

Some prior art reactors have relied upon arrangements that contain thereactions in generally adiabatic reactor zones and supply indirectcontact with a cooling medium between stages. The geometry ofintercooled reactors poses layout constraints that require largereactors and vast tube surfaces to achieve high heat transferefficiencies. In U.S. Pat. No. 4,696,799, Noe discloses an ammoniasynthesis converter having shell and tube interchangers for cooling thereaction gas streams leaving the catalyst beds with incoming reactantgases. In U.S. Pat. No. 5,869,011, Lee discloses a fixed bed reactorthat partitions a single stage catalyst bed into multiple heatinterchanged stages in a single reaction vessel.

In U.S. Pat. No. 6,171,570, Czuppon discloses maintaining asubstantially isothermal condition by boiling water on the shell side ofa shell-and-tube reactor with catalyst-filled tubes. Disclosedadvantages include lower energy consumption, lower purge rates, andhigher ammonia production rates. While overall catalyst efficiency canbe better than is found in reactors operating at adiabatic conditions,of course, the isothermal condition means that towards the outlet of thereactor, the reaction composition may still approach equilibrium productconcentrations, thus limiting further reaction. For an exothermicreaction such as ammonia production from hydrogen and nitrogen, theproduct ammonia concentration at the outlet end of an isothermal reactorcan be higher than the product concentration at the outlet of theadiabatic reactor given sufficient reaction time. This is true becausein the adiabatic reactor, with an exothermic reaction such as ammoniaproduction from hydrogen and nitrogen, temperatures increase along thereactor length and the equilibrium product concentration of ammonia islower at higher temperatures.

Adiabatic fixed bed reactors with interstage cooling have been used inthe prior art to provide successive conversion at lower and lowertemperatures to improve catalyst efficiency and improve yields. Inpractice, prior art reaction processes have been limited to two to fourstages in one common reactor vessel, with the major limitation being thecapital costs associated with interstage heat exchange equipment andmultiple reactor stages and/or vessels. In addition, inlet temperatureat each bed is necessarily lower than the outlet temperature, which iscloser to the equilibrium temperature. For example, in U.S. Pat. No.6,015,537, Gam discloses a reactor for the preparation of ammonia from asynthesis gas featuring multiple catalyst beds with intermediate coolingof partially converted synthesis gas between each catalyst individualbed.

In one commercially available prior art ammonia process, four catalystsbeds are provided with inter-cooling between each of the beds. The firstbed, and sometimes the second bed, can feature an iron-based magnetitecatalyst, followed by two or three beds which contain a ruthenium-basedcatalyst. Reactor temperature at the inlet of each catalyst bed is lowdue to the increasing temperature profiles in the adiabatic exothermicammonia synthesis reaction zones. The exothermic nature of the reaction,together with the adiabatic reactor bed design, do not allow thetemperature profile to maximize per-pass ammonia conversion, in turnleading to inefficient catalyst use. In such a system, larger amounts ofthe catalyst are necessary to achieve higher per pass ammoniaconversion.

Similarly, isothermal reactors have limitations in the production ofammonia. Synthesis of ammonia using an isothermal reactor generallyrequires separate external preheating of the feed gas. Additionally, aswith the staged adiabatic reactors, typical isothermal reactors haverelatively high catalyst requirements to obtain equivalent conversionrates.

Accordingly, there is a need in the art for a reactor design whichcontrols the temperature of exothermic reactions along the length of thereactor that effectively utilizes a temperature:conversion operatingcurve that follows the equilibrium curve with a negative temperatureoffset, and thus maintains a high reaction rate and catalyst efficiencythroughout the catalyst bed volume.

SUMMARY OF THE INVENTION

The present invention is directed to a conversion process and apparatusfor synthesis of ammonia, for example, wherein reaction temperaturealong the length of the reactor can be controlled, resulting indecreased catalyst volumes and increased conversion. The reactiontemperature can be reduced along the length of the reactor, to followthe equilibrium curve for the desired product such that the productconcentration increases along the length of the reactor but neverreaches the equilibrium product concentration at the current reactionmixture temperature.

The present invention provides, in one embodiment, a conversion processuseful for ammonia synthesis, comprising: (a) introducing a gaseousreactant-rich stream at a feed temperature into a heat exchange passageof a heat exchanging reaction zone to pre-heat the reactant-rich streamto an inlet temperature; (b) introducing the pre-heated reactant-richstream at the inlet temperature into a countercurrentcatalyst-containing reaction passage to exothermically convert thereactant gas to a product gas to form a product-enriched mixture of thereactant and product gases; (c) indirectly transferring heat from thereaction passage to the heat exchange passage at a rate effective tomaintain the mixture of gases below the equilibrium temperature; and (d)recovering an effluent from an outlet from the reaction passage at adischarge temperature enriched in the product gas.

The product-enriched mixture can have a product equilibriumconcentration that increases with decreasing temperature and a reactionrate coefficient that increases with increasing temperature. The heattransfer rate in a decreasing-temperature section of the reactionpassage can exceed the heat of reaction to lower the temperature of themixture gases to a discharge temperature. The reactant gas can include amixture of nitrogen and hydrogen and the product gas can compriseammonia. The catalyst can include a transition metal and in oneembodiment can include platinum group metal. The catalyst can includeruthenium on a carbon support and in one embodiment can include promotedruthenium on a heat stabilized graphitic carbon support. The heatexchanging reaction zone can include a shell and tube heat exchanger,the heat exchange passage can comprise a shell-side passage through theheat exchanger, and the reaction passage can include a plurality oftubes containing catalyst. The reaction passage can include a pluralityof alternating catalyst-containing zones and reaction-limited zones inseries. The reaction passage section can be maintained at a temperaturewithin 30° C. of an equilibrium temperature of the gas mixture. Thereaction passage can include an adiabatic initial zone adjacent an inletof the reaction passage wherein the heat of reaction exceeds the rate ofheat transfer whereby the temperature of the gas mixture rises. Thereaction limited zones can be non-reactive and can be free of catalyst.

The present invention provides, in another embodiment, a conversionprocess for the synthesis of ammonia, comprising: (a) introducing areactant-rich stream comprising hydrogen and nitrogen at a feedtemperature into a shell-side passage of a shell-and-tube heatexchanging reactor to pre-heat the reactant-rich stream to an inlettemperature; (b) introducing the pre-heated reactant-rich stream fromthe shell-side passage at the inlet temperature into a reaction zonecontaining a plurality of catalyst-containing tubes to convert thehydrogen and nitrogen to ammonia to form an ammonia-enriched mixture ofhydrogen, nitrogen and ammonia; (c) indirectly transferring heat fromthe tubes to the reactant-rich stream at a rate effective to maintainthe mixture in the tubes below equilibrium temperature, wherein the heattransfer rate in a decreasing-temperature section of the reaction zoneexceeds heat of reaction to lower the temperature of the mixture to adischarge temperature; and (d) recovering an effluent at the dischargetemperature from outlet ends of the tubes enriched in ammonia and leanin nitrogen and hydrogen.

The catalyst can include a transition metal and in one embodiment caninclude a platinum group metal. The catalyst can include ruthenium on acarbon support and in one embodiment can include promoted ruthenium on aheat stabilized graphitic carbon support. The tubes can include aninitial temperature-increasing zone adjacent an inlet of the reactionpassage wherein a heat of reaction exceeds the rate of heat transfer andthe temperature of the gas mixture is increasing. The tubes can includea series of alternating catalyst-containing zones and reaction-limitedzones. The mixture in the decreasing temperature section can bemaintained at a temperature within 30° C. of the equilibrium temperaturefor the reaction. The method can further include passing thereactant-rich stream through an upstream reactor comprising magnetitecatalyst, and supplying an effluent from the magnetite reactor in seriesas the reactant-rich stream introduced shell-side to the shell-and-tubeheat exchanging reactor. In one embodiment, the reactant-rich streamintroduced shell-side to the shell-and-tube heat exchanging reactor caninclude a purge gas stream from an ammonia synthesis loop. The catalysttubes can be vertical and the gas mixture can flow downward through thetubes. The tubes can include an initial temperature-increasing zoneadjacent an outlet of the reaction passage wherein a heat of reactionexceeds the rate of heat transfer and the temperature of the gas mixtureis increasing. The reaction limited zones can be non-reactive and can befree of catalyst. The decreasing temperature section can be maintainedover a range of temperatures that follow the ammonia concentration vs.temperature equilibrium curve for the reactor pressure, maintaining ahigh reaction rate and high catalyst efficiency throughout the reactor.

In another embodiment, the invention provides a converter for ammoniasynthesis, comprising: (a) means for introducing a gaseous reactant-richstream at a feed temperature into a heat exchange passage of a heatexchanging reaction zone to pre-heat the reactant-rich stream to aninlet temperature; (b) means for introducing the pre-heatedreactant-rich stream at the inlet temperature into a catalyst-containingreaction passage to exothermically convert the reactant gas to a productgas to form a product-enriched mixture of the reactant and product gaseshaving a product equilibrium concentration that increases withdecreasing temperature and a reaction rate coefficient that increaseswith increasing temperature; (c) means for indirectly transferring heatfrom the reaction passage to the heat exchange passage at a rateeffective to maintain the mixture of gases below the equilibriumtemperature, wherein the heat transfer rate in a decreasing-temperaturesection of the reaction passage exceeds heat of reaction to lower thetemperature of the mixture of gases to a discharge temperature; and (d)outlet means from the reaction passage for recovering an effluentenriched in the product gas at the discharge temperature.

The converter can include means to limit heat transfer from the reactionpassage to the heat exchanger passage near the catalyst tube outletsection. The reactant gas can comprise a mixture of nitrogen andhydrogen, and the product gas can comprise ammonia. The catalyst caninclude a transition metal, and in one embodiment can include a platinumgroup metal. The catalyst can include ruthenium on a carbon support andin one embodiment can include a promoted ruthenium on a heat-stabilizedgraphitic carbon support. The heat exchanging reaction zone can includea shell and tube heat exchanger. The heat exchange passage can include ashell-side passage through the heat exchanger and thecatalyst-containing passage can include a tube-side passage through aplurality of vertical parallel tubes containing the catalyst. Thereaction passage can include a generally adiabatic initial zone adjacentan inlet of the reaction passage wherein the heat of reaction exceedsthe rate of heat transfer and the temperature of the gas mixture rises.The reaction passage can also include a plurality of alternatingcatalyst-containing zones and reaction limited zones. The reactionlimited zones can be non-reactive and can be free of catalyst. Theconverter can further include means for maintaining the temperature ofthe gas mixture in the tubes within 30° C. of equilibrium temperaturefor the ammonia concentration as the gas mixture passes through thetubes.

In another embodiment, the invention provides a converter for ammoniasynthesis, comprising: (a) a shell-and-tube heat exchanging reactorcomprising a shell-side heat exchange passage and a reaction passagecomprising a plurality of catalyst containing tubes; (b) an inlet forintroducing a reactant-rich stream comprising hydrogen and nitrogen at afeed temperature into the heat exchange passage to pre-heat thereactant-rich stream; (c) an inlet for introducing the pre-heated streamto the reaction passage; (d) a series of alternating catalyst containingreaction zones and reaction limited zones in the tubes, to convert thehydrogen and nitrogen to ammonia to form an ammonia-enriched mixture ofhydrogen, nitrogen and ammonia; and (e) an outlet from the reactor forrecovering an effluent from the tubes at the discharge temperatureenriched in ammonia and lean in nitrogen and hydrogen.

The shell-side passage can include a plurality of baffles to direct aflow of the reactant rich stream across the reaction tubes. Spacingsbetween baffles in the shell-side passage can be variable, for example,with closer spacing of the baffles near the inlet end of the reactiontubes and increased spacing between adjacent baffles near the outlet endof the reaction tubes. The reaction tubes can include a series ofalternating catalyst-containing zones and reaction-limited zones. Thecatalyst can include a transition metal and in one embodiment caninclude platinum group metal. The catalyst can include ruthenium on acarbon support and in one embodiment can include promoted ruthenium on aheat stabilized graphitic carbon support. The insert devices can beselected from the group consisting of: screens and rod; screens and wiremesh; screens and twisted tape; metallic or ceramic structured packing;metallic or ceramic mesh pads; metal or ceramic foam; and structuredmetallic packing. The converter can further include an upstream reactorcomprising magnetite catalyst and a discharge operatively connected tothe converter inlet. The converter can further include a plurality ofthe shell-and-tube heat exchanging reactors with the converter inletsconnected to the discharge from the magnetite reactor in parallel flow.The reactant-rich stream can be introduced shell-side to theshell-and-tube heat exchanging reactor and in one embodiment can includea purge gas stream from an ammonia synthesis loop. The converter caninclude tube shields to limit heat transfer from the reaction passage tothe heat exchanger passage near the tube outlet section. The tubes canbe vertical and the gas mixture can flow downward through the tubes. Thetubes can include an adiabatic section adjacent an inlet of the reactionpassage wherein a heat of reaction exceeds the rate of heat transfer andthe temperature of the gas mixture rises.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic side sectional drawing of an ammonia converteraccording one embodiment of the invention.

FIG. 1A is an enlarged schematic side section view of another embodimentof the lower portion of the ammonia converter of FIG. 1.

FIG. 2 is a schematic side sectional view of a catalyst tube in theconverter of FIG. 1.

FIG. 3 is a schematic drawing of one embodiment of the insert deviceused in the catalyst tube of FIG. 2.

FIG. 4 is a schematic drawing of an alternate embodiment of the insertdevice used in the catalyst tube of FIG. 2.

FIG. 5 is a schematic drawing of an alternate embodiment of the insertdevice used in the catalyst tube of FIG. 2.

FIG. 6 is a schematic drawing of an alternate embodiment of the insertdevice used in the catalyst tube of FIG. 2.

FIG. 7 is a schematic drawing of a multiple-bed converter arrangement ofan alternate embodiment of the invention.

FIG. 8 is a process schematic of an ammonia synthesis loop with aconverter according to an alternate embodiment of the invention.

FIG. 9 is a graphical illustration comparing the ammonia conversion andtemperatures of staged adiabatic (inter-cooled) ammonia converters andisothermal ammonia converters of the prior art, with a converteraccording to an embodiment of the present invention.

FIG. 10 is a converter loadsheet illustrating the duty, shell-side andtube-side temperatures as a function of cumulative catalyst volumeaccording to an embodiment of an ammonia converter according to thepresent invention.

FIG. 11 shows cooling/heating curves illustrating shell-side andtube-side temperatures as a function of cumulative heat exchangedaccording to the embodiment of FIG. 10.

FIG. 12 is a converter loadsheet illustrating the ammonia concentrationas a function of tube-side temperature relative to the equilibrium curveaccording to the embodiment of FIGS. 10-11.

FIG. 13 is a graph of the overall heat transfer coefficient as afunction of cumulative catalyst volume according to the embodiment ofFIGS. 10-12.

DESCRIPTION OF THE INVENTION

The invention will be described by way of examples with reference toFIGS. 1-13 which are not to be construed as a limitation upon theapparatus elements and process steps of the invention. The presentinvention is directed to a synthesis reactor for exothermic reactionswhich can better approximate the equilibrium temperature/concentrationcurve for the conversion of reactants to products. Catalyst usage andreaction yields can be improved over existing isothermal and stagedadiabatic reactor processes. The advantages of the present invention areaccomplished by withdrawing heat from the catalyst-filled tubes of theconverter to maintain the reaction temperature within the bulk of thelength of the catalyst-filled tubes close to an optimum. Within thisapplication the terms reactor and converter can be used interchangeably.

Examples of commercially practiced exothermic-type gas phase catalyticreactions which can be practiced in reactor 100, as depicted in FIG. 1,include: synthesis reactions for producing ammonia, synthesis reactionfor producing methanol, shift reaction for producing CO and H₂,methanation, hydrocarbon oxidation for synthesizing maleic anhydride;other exothermic reactions mentioned above; and the like. The reactor ofthe present invention is particularly well suited for the synthesis ofammonia, which is used hereinafter as an example for the purpose ofillustration. Catalyst usage, reaction rates and total ammoniaconversion can be improved over the existing technology employingmulti-bed adiabatic catalyst beds and heat exchangers between the bedsto cool the gas feed.

To maintain equilibrium favoring the synthesis of products in exothermicreactions, heat is typically removed as the reaction progresses.Generally, more heat is removed at the beginning of the reaction than isremoved at the end of the reaction. The temperature profile of thecatalyst tube can be maintained whereby the highest temperatures aregenerated at the inlet of the catalyst tube (i.e. near the top end ofthe tube with downward gas flow) and the temperatures at the outlet ofthe catalyst tube can be maintained at a minimum effective catalysttemperature to ensure high conversion rates. The rate of conversion canbe greater near the inlet end of the catalyst tube as the concentrationof ammonia can be low at this point of the reaction and the hightemperature ensures favorable kinetics. The rate of conversion can belower near the outlet end of the catalyst tube and more catalyst volumeis required per unit mass of product generated. The rate of heatgeneration can be relatively low when compared to the heat transferpotential due to high surface area which becomes available as increasedcatalyst is loaded at the lower end of the catalyst tube. To preventovercooling of the reaction gas, part of the heat transfer surface areacan be to be blocked so that heat transfer is inhibited in the blockedarea.

Referring to FIGS. 1 and 2, there is provided a reactor 100 consistingof a shell 101, baffles 106, internal catalyst tubes 109, and tubeshields 116. As shown in FIG. 1A, a number of flow blocker devices (116)can be provided throughout the reactor. The vertical height of theshield 116 can be selected to approximate the catalyst temperatureprofile as close to the maximum reaction rate profile as possible, byreducing the catalyst tube area exposed to and cooled by the inlet feedstream. The blocker device can be formed by using two consecutivespacers or baffles on the shell side. The feed stream enters reactor 100through inlet 102, located near the bottom of the reactor shell andflows upward through a series of horizontal baffles 106. The feed streamcan be preheated through indirect heat exchange with the catalyst tubes109 located vertically inside the reactor shell 101. Optionally, thefeed stream can be preheated upstream of the reactor inlet 102. Ifdesired, additional feed inlets, shown as 102A, 102B and 102C, can beadded to the reactor shell 101. The feedstock temperature can becontrolled by installing a bypass downstream from a preheater (notshown), whereby cold gas can be mixed with a portion of the preheatedgas to obtain an optimum inlet temperature. The baffles can, if desired,have variable spacing between adjacent individual baffles to helpgradually increase heat transfer on the approach to the inlet end 110 ofthe catalyst tube 109. For example, the spacing between adjacent bafflescan be less near the top of the reactor than at the bottom of thereactor. The catalyst tubes 109 are desirably supported on a tubesheet104, in a conventional manner such as by welding the tubes 109 to thetubesheet 104, or any other method known in the art. Tubesheet 104separates the tube side (reaction product) from the shell side (feedstream) of the reactor 100.

The reactor 100 can include a diffuser plate (not shown) located at theupper portion of the reactor near the inlet to the catalyst tubes toensure proper mixing of the feed gases. A bypass inlet 111 can bepositioned at the top of the reactor 100 and can provide a preheatedfeed stream if desired. Bypass inlet 111 can be used to control thetemperature of the feed gas at the inlet to the catalyst tubes 109. Adiffuser plate (not shown) can be located near the bypass inlet 111 toensure proper mixing with the feed supplied to inlet 102. The feedstream enters catalyst tubes 109 through catalyst tube inlets 110located at the uppermost portion of the vertical catalyst tubes 109.Generally, the feed stream flows countercurrent to the direction of theflow through the catalyst tubes 109, i.e. the reactant gas flows upwardin the shell-side of the reactor and downward through the catalysttubes. The converted effluent exits the catalyst tubes 109 at thecatalyst tube outlet 113, located at the lower most portion of thecatalyst tube 109 and the lower portion of the reactor 100. A screen 114located below the tubesheet 104 prevents loss of catalyst to the outletzone of the reactor. The product effluent can exit the reactor 100through reactor outlet 112.

Referring now to FIG. 2, the cross sectional detail of the catalyst tube109 located in reactor 100 is provided. The catalyst tube is desirablybetween 25 and 75 mm in diameter, more desirably between 38 and 63 mm indiameter. The catalyst tube length is desirably between 3 and 8 m, moredesirably between 4 and 6 m. The tube consists of alternating reactivezones 122 and reaction limited zones 124. The reactive zones 122 canconsist of high activity catalyst and the reaction limited zones 124 canconsist of non-reactive spacer insert devices designed to separate thecatalyst zones and transfer a portion of the heat of reaction to theshell side of the reactor. The reaction limited zones can benon-reactive and in one embodiment the reaction limited zones can befree of catalyst. Baffles 106 can be located within the interior of thereactor shell 101 to guide the fluid flow and facilitate heat transferbetween the feed stream and the catalyst tubes 109. The catalyst tube109 can be supported at the base by tubesheet 104. Catalyst screen 114,located at the bottom of the catalyst tube, facilitates catalystretention at the bottom of tube 109. The catalyst tube can also consistof a screen or cap 110 located at the uppermost portion of the verticaltube.

Typically, the converter can be sized and designed based upon theoperating conditions of the catalyst as the activity of the catalystdecreases. To achieve good operability over the entire run, including atthe start of the run when the catalyst activity is high, cooling in thelower portion of the converter can be minimized or eliminated. One ormore bypass inlets 102A, 102B and 102C (see FIG. 1) can be used tobypass the cooling section for the corresponding lower portion of thecatalyst tubes 109. The bypass feature can be used depending onoperating conditions and activity of the catalyst and can be used toavoid overcooling of the catalyst when the operating conditions aredifferent than those used when sizing the converter. Overcooling thecatalyst can cause the reaction to run at a temperature too low toobtain optimum conversion of the reactants at the outlet end of thecatalyst tube.

Because the catalyst tubes can be mounted in a vertical arrangement, thescreen 114 located at the bottom plays an important role in retainingcatalyst particles in the tube 109. Screens such as Johnson Vee-Wirescreens supplied by UOP can be well suited for use in the convertersystem. The insert devices 124 can comprise a variety of shapes as shownin FIGS. 3-6. Alternatively, twisted tape inserts, such as for example,turbulators manufactured by Brown Fintube, HiTRAN® Matrix Elementmanufactured by CalGavin Co., or static mixers available from a varietyof manufacturers and welded to a Johnson screen, can be used as theinsert device 124 for the non-reactive or reduced activity zone. Theinsert devices 124 can be constructed from a variety of materials,selected based upon the heat transfer properties of the chosen material.The insert devices 124 can be fixed in place within the catalyst tube109, or optionally, they can be placed between layers of catalyst 122whereby the insert devices 124 can move cooperatively with the catalyst122, as catalyst settling may occur during normal operation. Dependingon the type of insert device 124, it can be desirable to weld the insertdevices to a rod to maintain reactive zones of a certain catalyst volumeand a certain location, thereby preventing the catalyst zones 122 frommoving within the tube 109. It is anticipated that the catalyst volumeof a given reactive zone can vary by up to 10%, from the beginning ofthe reaction until the end of the reaction, as catalyst settling canoccur during operation of the reactor.

A variety of materials, such as for example, insert elements and screenscan be used as the insert devices 124 for the non-reactive or reducedactivity zone within the catalyst tube 109, including: screens and rod;screens and wire mesh; screens and twisted tape; metal foam, staticmixing type inserts; and the like. The screen can be attached with aseal (e.g. a leaf type seal) about the circumference of the insert 124.A means to inhibit catalyst from passing through, or lodging between,the screen and tube wall can also be provided. The means of inhibitingcatalyst from passing through or around the non-reactive or reducedactivity zone can compensate for typical variations in reactor tubediameter and tube cross-sectional ovality. Ideally, the insert devices124 can perform multiple functions including, but not limited to,providing a non-reactive or reduced activity zone within the catalysttube 109, enhancing heat transfer between the shell and tube side of thereactor, and enhancing mixing of fluids on the tube side of the reactor.Twisted tape inserts can be used as the insert device 124, desirablyhaving a length coextensive with the desired length of the reducedreactivity or non-reactive zone. The twisted tape can have an outerdiameter that is approximately equal to the inner diameter of the tube,such that a ratio of the tube diameter to the packing diameter isapproximately 1. The insert can also be one or more stacked pieces ofmetallic structured packing, such as Sulzer Type DX, Type EX, or TypeDXM/DYM laboratory packing made by Sulzer Chemtech designed to fitclosely within the inner diameter of the reactor tubes. One or moreshorter length pieces of the structured packing can be stackedend-to-end in various multiples and combinations to form a number ofdifferent length non-reactive or reduced activity zones interspersedwith different length zones of the reactive catalyst as required toachieve the desired configuration. Similarly, the non-reactive insertcan be one or more metallic mesh pads, such as Hyperfil® High EfficiencyColumn Packing, a knitted mesh distillation packing made by EnhancedSeparation Technologies, LLC. The metallic mesh pads can be designed tofit closely within the inner diameter of the reactor tubes and can bestacked end-to end when multiple pads are used to create a singlecontiguous non-reactive zone of the desired length. This can reduce thecomplexity of the insert fabrication.

The non-reactive insert can be ceramic foam in another embodiment. Theceramic foam can be made by filling voids in an organic sponge with afluidized ceramic precursor and burning the substrate away to form theceramic foam. Advantageously, the ceramic foam can be cut into shorterlengths, and can be stacked end-to-end in each tube to achieve a limitedor non-reactive zone of the desired length.

A variety of catalysts can be used with the present invention, includingtraditional ammonia synthesis magnetite-based catalysts. Desirably, aruthenium based catalyst can be used in the reactor catalyst tubes. Theruthenium catalyst can have an effective diameter of between 1.5 and 2.0mm, and can be formed in a variety of shapes, on a carbon-based supportmaterial. The catalyst can be a promoted ruthenium on a heat stabilizedcarbon support. The ruthenium based catalyst can exhibit an activity upto 20 times greater than magnetite, and performance can be maintained athigh ammonia concentrations and over a wide range ofhydrogen-to-nitrogen ratios. Additionally, the ruthenium catalyst canallow the ammonia synthesis to be conducted at lower pressures thanneeded for synthesis performed with magnetite catalyst.

Replacement of the catalyst can be accomplished in a variety of ways. Asillustrated in FIG. 1, the reactor 100 can feature a flange 115,allowing the separation and removal of the top portion of shell 101 toaccess the catalyst tubes 109. The tubesheet 104 can be fixed in placeand catalyst can be removed from the tubes via vacuum means.Alternatively, the reactor 100 can be arranged with channel 113 locatedat the top of the reactor shell 100. With such an arrangement, thecatalyst tubes 109 and tubesheet 104 are then supported at the top ofthe reactor shell 101 and the entire bundle 104 and 109 can be removedfor replacement of catalyst.

In another embodiment of the present invention, shown in FIG. 7,multiple reactors can be used in parallel. Vessel 200 features tworeactors, 201 a and 201 b, each having vertical catalyst tubes 216 a and216 b, arranged in parallel. A first feed stream start up gas 202 isintroduced to the parallel reactors 201 a and 201 b at inlets 204 a and204 b respectively, each located at the top of their respectivereactors. The start up gas 202 can be preheated as desired. Inlets 204 aand 204 b can be located above the tops of the catalyst tubes 216 a and216 b located within the reactors 201 a and 201 b, respectively. Asecond feed stream 205 from a synthesis loop using lower activitycatalysts, such as for example, magnetite, can be introduced to thereactors via inlets 206 a and 206 b, located near the base of thecatalyst tubes 216 a and 216 b. The reactors 201 a and 201 b can includea series of baffles 212 a and 212 b through which second feed stream 205passes and undergoes indirect heat exchange with the catalyst tubes 216a and 216 b. The baffles 212 a and 212 b can optionally be spaced closerat the upper most portion of the reactor, as illustrated in FIG. 7, andcan thereby increase heat exchange between the second feed stream 205and the upper portion of the catalyst tubes 216 a and 216 b. The firstheated start-up feed stream 202 and heated second feed stream 205 enterthe catalyst tubes 216 a and 216 b at inlets 214 a and 214 b,respectively. The catalyst tubes 216 a and 216 b can feature alternatingreactive and non-reactive zones, as previously described and illustratedin FIG. 2. Product effluent exits the catalyst tubes 216 a and 216 b,passes through catalyst screens 210 a and 210 b, and exits reactors 201a and 201 b via outlet 220 a and 220 b. Product effluent from reactors201 a and 201 b can be combined to form a product effluent 222, whichcan be introduced into a series of heat exchangers (not shown) to coolthe product and indirectly heat the feed gas, or can optionally be usedfor other heat recovery processes. In a similar fashion, the convertercan be used in conjunction with an upstream fixed bed magnetite reactor(not shown). Generally adiabatic magnetite bed(s) and core heatexchangers can also be housed in the vessel 200 as indicated in FIG. 7.

A process schematic of an example of an ammonia synthesis loop having aconverter according to another embodiment of the invention is shown inFIG. 8. Synthesis feed gas 302 enters compressor 304, and in the exampleshown, the exiting gas has a temperature of 70° C. (158° F.) and apressure of 9.43 MPa. If desired, a cryogenic purification step can beused to provide high purity makeup gas to the ammonia synthesis loop.The purification step desirably removes excess nitrogen, along withmethane and argon, from the feed gas. The compressed feed gas flows viastream 306 to first cross exchanger 308, where it is heated to 234° C.(453° F.), and then via stream 310 to second cross exchanger 312, wherethe feed gas is further heated to 357° C. (675° F.). The heated feedflows via path 314 to a radial flow magnetite converter 316, having amagnetite catalyst bed and enters the converter via inlet 318.Desirably, magnetite converter 316 can have more than one catalyst bed.After cooling in magnetite heat exchanger 317, effluent from themagnetite converter 316 exits via outlet 320, at a temperature in thisexample of 392° C. (737° F.) and a pressure of 9.18 MPa. The effluent iscooled via second cross exchanger 312 to a temperature of 266° C. (511°F.). The effluent is further cooled via first heat exchanger 322 to atemperature of 217° C. (423° F.) and a pressure of 9.11 MPa.

The cooled effluent 324 enters converter 326 via inlet 328, flowsthrough a series of baffles in indirect heat exchange with the catalysttubes 330, flows to the inlet end 331 of the catalyst tube 330, andexits catalyst tubes 330 and reactor 326 via outlet 334. The catalysttubes 330 can consist of alternating layers of catalyst and inertspacers, as previously discussed and shown in FIG. 2, thereby creatingreactive and non-reactive zones in each catalyst tube 330. The catalystcan be a ruthenium based catalyst, desirably promoted ruthenium on aheat stabilized graphitic carbon support, although other similar highactivity catalysts or support materials can be used. The ammoniaeffluent from the converter 326 has a temperature of 373° C. (703° F.)and a pressure of 8.90 MPa, in this example. Total pressure drop for theconverter is desirably less than 0.2 MPa.

The hot ammonia effluent exiting the converter 326 is available for heatrecovery and flows via line 338 to magnetite converter heat exchanger317 where it is heated to a temperature of 448° C. (839° F.). Theammonia stream exits the magnetite converter via 343 and passes throughsecond heat exchanger 344 and exits the second heat exchanger 344 vialine 345 where the ammonia stream has cooled to a temperature of 261° C.(502° F.). Stream 345 flows into first cross exchanger 308 for furthercooling and into line 348 where the cooled ammonia stream has atemperature of 87° C. (188° F.) and a pressure of 8.76 MPa, and flowsinto chilled water cooler 350. The cooled ammonia enters a conventionalrefrigeration unit 354 via line 352, where the reactor effluent can besplit into three streams. Stream 356 recycles a mixture of ammonia andhydrogen to the compressor 304 via line 362 where it is combined withsynthesis gas 302, or alternatively, separates out ammonia and hydrogenvia conventional ammonia recovery unit 363 and provides a hydrogenstream 364 for desulfurization, for example. The purge gas can befurther processed in a membrane hydrogen-recovery unit, which canrecover up to 90% of the hydrogen for recycle. Waste gas from themembrane separator can be combined with other purge gas streamsgenerated by the process and used as supplemental fuel. Stream 358provides a mixture which can either be recycled to a purification unit(not shown) for hydrogen recovery or purged to fuel for the reactors(not shown). Stream 360 provides a purified ammonia product stream.

The converter can be designed to operate with one of two goals in mind:(a) minimizing the catalyst usage or (b) maximizing the ammoniaconversion. In the first scenario, the converter can operate using alower catalyst volume, while still obtaining high ammonia yields. Thus,the converter can operate with lower catalyst costs than similarcatalyst systems employing adiabatic reactor beds. In the secondscenario, the reactor can be operated using a larger catalyst volumethan the first scenario, thereby obtaining greater ammonia conversion atreduced catalyst efficiency.

The converter can be well suited for debottlenecking existing plants,thereby increasing production and/or productivity. The converter can beinstalled as an “add-on” feature to an existing plant for recovery andconversion of purge gas streams. Because the converter can be deployedin a retrofit application as an “add-on” converter to supplement ammoniaproduction from one or more existing ammonia synthesis plants, baseproduction can be unaffected during installation. In addition, verylittle plant downtime would be experienced during the retrofitturnaround shutdown to make the required tie-ins to complete theretrofit installation.

The converter demonstrates an ability to provide higher conversion offeed by closely approximating the equilibrium line for the reaction.FIG. 9 illustrates a comparison between theoretical temperature:ammoniaconversion curves for the reaction equilibrium line 400, a three-stageadiabatic inter-cooled reactor 402, an isothermal reactor 404, and areactor 406. The theoretical reaction equilibrium line 400, as shown inFIG. 9, demonstrates that with an unlimited residence time, conversionincreases with decreasing temperature. The three-stage adiabatic reactor402 employs a series of three inter-coolers to reduce the temperaturebetween stages in the reactor and increase the ammonia conversion. Theheating and cooling cycles are discernable as there is no conversionduring the cooling step. Conversion and rate of reaction using anisothermal reactor 404 is limited by the temperature at which thereaction is run. As shown in FIG. 9, the reactor 406 provides highconversion efficiency as the reactor provides increased conversion astemperature decreases, following that of the calculated reactionequilibrium.

Catalyst requirements can be significantly reduced and ammoniaconversion for the reactor can be improved when compared with the stagedadiabatic and isothermal reactors. For example, to achieve equivalentsingle-pass conversions using a high-activity ammonia catalyst, thecatalyst requirements for the present invention reactor, in someembodiments, can be approximately 30% less than that required for thestaged adiabatic reactor. Whereas typical ammonia conversion processesin staged adiabatic reactors exhibit single-pass conversions yieldingammonia concentrations in the reactor effluent of approximately 20.6 mol%, the reactor of the present invention is capable of achieving singlepass conversions yielding ammonia concentrations in the reactor effluentgreater than 22 mol %, and in some embodiments greater than 22.8 mol %when using an excess catalyst load of a high-activity ruthenium-basedcatalyst.

The reactor can have between 200 and 10,000 catalyst tubes, desirablybetween 1,500 and 2,500 catalyst tubes. Each reactor can contain betweenapproximately 30 and 35 m³ of ruthenium-based catalyst, loaded incatalyst tubes interspaced with non-reactive inserts as previouslydescribed and shown in FIG. 2. FIG. 10 shows the temperature change forboth the tube and shell side of the reactor per cubic meter of catalyst,as well as total heat produced at a particular volume of catalyst in asimulation example, for one embodiment of the invention. As shown, thefirst 1 to 6 m³ of catalyst is assumed to be a generally adiabatic zone,where the reaction can proceed without transfer of all or a portion ofthe heat generated by the exothermic reaction. In this example, shellside temperature in the generally adiabatic zone remains essentiallyconstant while the tube side temperature increases to a maximum ofbetween 416-427° C. (780-800° F.). Heat transfer begins after thegenerally adiabatic zone, and totals between 135 and 150 GJ (130 and 140MMBtu).

FIG. 11 shows a heating and cooling curve for the tube and shell sidesof the reactor simulation of FIG. 10 as a function of cumulative heatexchanged. Downstream from the generally adiabatic zone, the temperatureof the tube side decreases from approximately 416-427° C. (780-800° F.)to approximately 377° C. (710° F.). The linear relationship between thetube side and shell side temperatures demonstrates good heat exchange.

FIG. 12 shows ammonia concentration as a function of tube sidetemperature relative to the equilibrium for the reaction for thesimulation of FIGS. 10-11. Tube side temperature in the generallyadiabatic zone shows a steady increase in temperature to approximately421° C. (790° F.), at which point the operating curve begins to parallelthe equilibrium curve for the reaction with the reaction temperaturebeing maintained within 4-15° C. (40-60° F.) of the equilibriumtemperature for the same ammonia concentration and the ammonia contentof the reactant gases at 4-6 mol % below the equilibrium ammonia contentat the same reaction temperature. The decrease in temperature isobtained because the indirect heat exchange between the shell side feedand the catalyst tube exceeds the heat of reaction by a controlledamount. The composition of the feed gas to the reactor in this exampleis approximately 55.9% H₂, 29.0% N₂, 3.9% CH₄, 1.9% Ar, and 9.2% NH₃.Feed to the shell side inlet has an initial temperature of 217° C. (423°F.). Product effluent has a composition of approximately 46.1% H₂, 26.9%N₂, 4.3% CH₄, 2.1% Ar, and 20.6% NH₃. Total conversion as shown in FIG.12 is in excess of 20% at an output temperature of between 371-377° C.(700-710° F.), and a pressure of approximately 8.85 MPa.

The product of the heat transfer coefficient times the surface area (UA)necessary to keep the reactor running parallel to the equilibrium curveof FIG. 12 is shown in FIG. 13. The required UA is greatest at theinitial part of the reaction, after the generally adiabatic zone. The UArequired at the inlet of the catalyst zone is approximately 14 timesgreater than that required at the outlet, and can be achieved throughthe use of extended surfaces, varied surface area, non-reactive spacers,baffles, or the like.

The reactor can achieve increased ammonia conversion and catalystefficiency by controlling temperature in the catalyst tube. Increasedheat transfer between the catalyst tube and the effluent on the shellside of the reactor can be achieved through the use of non-reactiveinserts placed inside the interior of the catalyst tubes, as shown inFIG. 2. These inserts can increase the surface area available for heattransfer for an equivalent volume of catalyst. Alternatively, portionsof the feed can be introduced at the bottom of the reactor, andpreheated by the baffles. One or more additional portion of the feed canbe introduced to the reactor at a point or various points below theinlet of the catalyst tube, desirably at locations below the generallyadiabatic zone. The catalyst tubes can include fins at the top portionof the catalyst tube to increase surface area for heat transfer. Thiscan be desirable when the shell side heat transfer coefficient isdominant. Alternatively, the baffle spacing between adjacent baffles canbe varied such that the spacing decreases toward the top of the catalysttube, i.e. the spacing between baffles is greater at the bottom of thereactor than at the top of the reactor. Catalyst tube shapes can also bevaried to increase catalyst volume at different portions of the tube.For example, the catalyst tube can use a slightly conical shape, ratherthan a cylindrical shape, providing a greater catalyst volume per tubearea at the bottom of the catalyst tube than at the top of the catalysttube.

The converter in one embodiment of the present invention can be used totreat purge gas for incremental increases in the production of ammonia.In ammonia synthesis loops, a purge stream can be employed to removeinert gases that may accumulate in the loop. The purge stream typicallycontains ammonia, H₂, N₂ and inert gases (CH₄, Ar and He). The purgestream can be processed to remove ammonia and/or hydrogen, and used as afuel gas. In the case of the present converter, the purge flow rate canbe increased significantly to reduce the amount of inert gases at themain inlet to the converter. In the practice of prior art processes,increasing the purge rate can be uneconomical because it can lead to thewasting of high pressure syngas and the compression power of the productincreases. The converter can convert the bulk of the H₂ and N₂ in thepurge gas to ammonia, and the net purge gas flow rate can be maintained.The converter can be used for debottlenecking, or providing incrementalproduction for a number of ammonia plants located at the same location,as the purge stream from each plant can be supplied to, and processedin, a single converter.

The invention is described above in reference to specific examples andembodiments. The metes and bounds of the invention are not to be limitedby the foregoing disclosure, which is illustrative only, but should bedetermined in accordance with the full scope and spirit of the appendedclaims. Various modifications will be apparent to those skilled in theart in view of the description and examples. It is intended that allsuch variations within the scope and spirit of the appended claims beembraced thereby.

1. A conversion process for ammonia synthesis, comprising: introducing agaseous reactant-rich stream at a feed temperature into a heat exchangepassage of a heat exchanging reaction zone to pre-heat the reactant-richstream to an inlet temperature; introducing the pre-heated reactant-richstream at the inlet temperature into a countercurrentcatalyst-containing reaction passage to exothermically convert thereactant gas to a product gas to form a product-enriched mixture of thereactant and product gases; indirectly transferring heat from thereaction passage to the heat exchange passage at a rate effective tomaintain the mixture of gases below the equilibrium temperature; andrecovering an effluent from an outlet from the reaction passage at adischarge temperature enriched in the product gas.
 2. The method ofclaim 1, wherein the product-enriched mixture has a product equilibriumconcentration that increases with decreasing temperature and a reactionrate coefficient that increases with increasing temperature.
 3. Themethod of claim 1, wherein the heat transfer rate in adecreasing-temperature section of the reaction passage exceeds heat ofreaction to lower the temperature of the mixture gases to a dischargetemperature.
 4. The method of claim 1, wherein the reactant gascomprises a mixture of nitrogen and hydrogen and the product gascomprises ammonia.
 5. The method of claim 1, wherein the catalystcomprises a transition metal.
 6. The method of claim 1, wherein thecatalyst comprises ruthenium on a carbon support.
 7. The method of claim1, wherein the heat exchanging reaction zone comprises a shell and tubeheat exchanger, the heat exchange passage comprises a shell-side passagethrough the heat exchanger, and the reaction passage comprises aplurality of tubes containing catalyst.
 8. The method of claim 1,wherein the reaction passage comprises a plurality of alternatingcatalyst-containing zones and reaction-limited zones in series.
 9. Themethod of claim 1, wherein the reaction passage section is maintained ata temperature within 30° C. of an equilibrium temperature of the gasmixture.
 10. A conversion process for ammonia synthesis, comprising:introducing a reactant-rich stream comprising hydrogen and nitrogen at afeed temperature into a shell-side passage of a shell-and-tube heatexchanging reactor to pre-heat the reactant-rich stream to an inlettemperature; introducing the pre-heated reactant-rich stream from theshell-side passage at the inlet temperature into a reaction zonecontaining a plurality of catalyst-containing tubes to convert thehydrogen and nitrogen to ammonia to form an ammonia-enriched mixture ofhydrogen, nitrogen and ammonia; indirectly transferring heat from thetubes to the reactant-rich stream at a rate effective to maintain themixture in the tubes below equilibrium temperature, wherein the heattransfer rate in a decreasing-temperature section of the reaction zoneexceeds heat of reaction to lower the temperature of the mixture to adischarge temperature; and recovering an effluent at the dischargetemperature from outlet ends of the tubes enriched in ammonia and leanin nitrogen and hydrogen.
 11. The method of claim 10, wherein thecatalyst comprises a platinum group metal.
 12. The method of claim 10,wherein the catalyst comprises ruthenium on a carbon support.
 13. Themethod of claim 10, wherein the tubes comprise an initialtemperature-increasing zone adjacent an inlet of the reaction passagewherein a heat of reaction exceeds the rate of heat transfer and thetemperature of the gas mixture is increasing.
 14. The method of claim10, wherein the tubes comprise a series of alternatingcatalyst-containing zones and reaction-limited zones.
 15. The method ofclaim 10, wherein the mixture in the decreasing temperature section ismaintained at a temperature within 30° C. of the equilibrium temperaturefor the reaction.
 16. The method of claim 10 further comprising passingthe reactant-rich stream through an upstream reactor comprisingmagnetite catalyst, and supplying an effluent from the magnetite reactorin series as the reactant-rich stream introduced shell-side to theshell-and-tube heat exchanging reactor.
 17. The method of claim 10,wherein the reactant-rich stream introduced shell-side to theshell-and-tube heat exchanging reactor comprises a purge gas stream froman ammonia synthesis loop.
 18. A converter for ammonia synthesis,comprising: means for introducing a gaseous reactant-rich stream at afeed temperature into a heat exchange passage of a heat exchangingreaction zone to pre-heat the reactant-rich stream to an inlettemperature; means for introducing the pre-heated reactant-rich streamat the inlet temperature into a catalyst-containing reaction passage toexothermically convert the reactant gas to a product gas to form aproduct-enriched mixture of the reactant and product gases having aproduct equilibrium concentration that increases with decreasingtemperature and a reaction rate coefficient that increases withincreasing temperature; means for indirectly transferring heat from thereaction passage to the heat exchange passage at a rate effective tomaintain the mixture of gases below the equilibrium temperature, whereinthe heat transfer rate in a decreasing-temperature section of thereaction passage exceeds heat of reaction to lower the temperature ofthe mixture of gases to a discharge temperature; and outlet means fromthe reaction passage for recovering an effluent enriched in the productgas at the discharge temperature.
 19. A converter for ammonia synthesis,comprising: a shell-and-tube heat exchanging reactor comprising ashell-side heat exchange passage and a reaction passage comprising aplurality of catalyst containing tubes; an inlet for introducing areactant-rich stream comprising hydrogen and nitrogen at a feedtemperature into the heat exchange passage to pre-heat the reactant-richstream; an inlet for introducing the pre-heated stream to the reactionpassage; a series of alternating catalyst containing reaction zones andreaction limited zones in the tubes, to convert the hydrogen andnitrogen to ammonia to form an ammonia-enriched mixture of hydrogen,nitrogen and ammonia; and an outlet from the reactor for recovering aneffluent from the tubes at the discharge temperature enriched in ammoniaand lean in nitrogen and hydrogen.
 20. The converter of claim 19,wherein the shell-side passage comprises a plurality of baffles todirect the flow of the reactant rich stream across the reaction tubes.21. The converter of claim 19, wherein spacings between baffles in theshell-side passage is variable with closer spacing of the baffles nearthe inlet end of the reaction tubes and increased spacing betweenadjacent baffles near the outlet end of the reaction tubes.
 22. Theconverter of claim 19, wherein the reaction tubes comprise a series ofalternating catalyst-containing zones and reaction-limited zones. 23.The converter of claim 19, wherein catalyst containing tubes compriseruthenium on a carbon support.
 24. The converter of claim 19, whereinthe insert devices is selected from the group consisting of screens androd; screens and wire mesh; screens and twisted tape; metallic orceramic structured packing; metallic or ceramic mesh pads; metal orceramic foam; and structured metallic packing.
 25. The converter ofclaim 19, further comprising an upstream reactor comprising magnetitecatalyst and a discharge operatively coupled to the converter inlet. 26.The converter of claim 19, further comprising a plurality of theshell-and-tube heat exchanging reactors for receiving the effluent fromthe magnetite reactor in parallel flow.
 27. The converter of claim 19,wherein the reactant-rich stream introduced shell-side to theshell-and-tube heat exchanging reactor comprises a purge gas stream froman ammonia synthesis loop.